Counter-Current Fluidized Bed Reactor for the Dehydrogenation of Olefins

ABSTRACT

A process and apparatus for the dehydrogenation of paraffins is presented. The process utilizes a reactor that includes a slower flow of catalyst through the reactor, with a counter current flow of gas through the catalyst bed. The catalyst is regenerated and distributed over the top of the catalyst bed, and travels through the bed with the aid of reactor internals to limit backmixing of the catalyst.

FIELD OF THE INVENTION

The field of the invention is production of light olefins. Inparticular, the invention relates to the dehydrogenation of paraffins inthe C3 to C5 range.

BACKGROUND OF THE INVENTION

Ethylene and propylene are light olefin hydrocarbons with two or threeatoms per molecule, respectively, are important chemicals for use in theproduction of other useful materials, such as polyethylene andpolypropylene. Polyethylene and polypropylene are two of the most commonplastics found in use today and have a wide variety of uses for both asa material fabrication and as a material for packaging. Other uses forethylene and propylene include the production of vinyl chloride,ethylene oxide, ethylbenzene and alcohol. Steam cracking or pyrolysis ofhydrocarbons produces essentially all of the ethylene and propylene.Hydrocarbons used as feedstock for light olefin production includenatural gas, petroleum liquids, and carbonaceous materials includingcoal, recycled plastics or any organic material.

A light olefin plant is a very complex combination of reaction and gasrecovery systems. The feedstock is charged to a cracking zone in thepresence of steam at effective thermal conditions to produce a pyrolysisreactor effluent gas mixture. The pyrolysis reactor effluent gas mixtureis stabilized and separated into purified components through a sequenceof cryogenic and conventional fractionation steps. A typical lightolefin plant includes an ethylene separation section containing bothcryogenic and conventional fractionation steps to recover an ethyleneproduct with a purity exceeding 99.5% ethylene. Propylene and heavierhydrocarbons are separated from the ethylene and recovered in a separatesection, or separate fractionation column.

Modification of the process can save energy, and equipment which arevery expensive, while increasing the overall yields of product.

SUMMARY OF THE INVENTION

The present invention is a process for the dehydrogenation ofhydrocarbons. In particular, the present invention is for a gas phaseconversion of alkanes in the C2 to C5 range to olefins. The processincludes using a vessel having a large diameter for flowing a catalystdownward through a catalyst bed. The process includes flowing a gasphase hydrocarbon process stream upward through the catalyst bed todehydrogenate the hydrocarbons as the process stream contacts thecatalyst. The catalyst is passed to the reactor to a region above thecatalyst bed at a temperature of at least 600° C. The process stream ispassed to the reactor at a temperature less than 600° C.

The process involves the counter-current flow of catalyst with respectto the process stream, and the process stream picks up heat from thecatalyst as it contacts the catalyst during the endothermic reaction ofdehydrogenation. A temperature gradient is formed within the catalystbed with the top of the catalyst bed at or near the highest temperatureand the bottom of the catalyst bed at or near the lowest temperature. Itis preferred to have a temperature difference between the process streaminlet temperature and the catalyst inlet temperature, wherein thecatalyst is at least 50° C. hotter than the process stream temperature.

Other objects, advantages and applications of the present invention willbecome apparent to those skilled in the art from the following detaileddescription and drawings.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a plot of a process stream temperature for a typical processutilizing multiple reactors;

FIG. 2 is a schematic of the design for a dehydrogenation systemincluding reactor and regenerator system;

FIG. 3 is a plot of the process stream temperature for thecounter-current flow;

FIG. 4 is a diagram of the reactor;

FIG. 5 shows one example of the reactor internals; and

FIG. 6 shows a second example of a part of the reactor internals.

DETAILED DESCRIPTION OF THE INVENTION

Currently, the production of light olefins is primarily from the normalsources of light olefins that are produced through the cracking processof naphtha and heavier hydrocarbons, and through the process of crackingheavier olefins. Light olefins are subsequently separated out from aproduct stream comprising ethylene and propylene. There is a growing gapbetween the production of light olefins and the demand for these polymerbuilding blocks. The demand is being met through dedicated processesthat use light paraffinic feedstocks, and directly convert the paraffinsto olefins through dehydrogenation. One example of a preferred feedstockis propane or an LPG feed. This can be directly dehydrogenated andovercomes drawbacks of other methods of propylene production, such asmethanol to olefins and the cracking of heavier hydrocarbons.

The production of light olefins using a process for the directconversion of a paraffin feedstream to the olefin analog utilizes anoble metal catalyst. The challenges in dehydrogenation technologyinclude the reaction conditions, such as pressure and temperature, thatfavorably shift the dehydrogenation equilibrium towards olefins, and thelarge amount of heat required to drive the reaction, while minimizingundesirable side reactions, such as non-selective thermal conversion.The dehydrogenation process is endothermic, and currently the processutilizes a plurality of reactor beds with interstage heating between thereactor beds. As shown in FIG. 1, the process comprises heating the feedto each reactor, where the process stream cools due to theendothermicity.

In one operation (A), the process stream as it exits one reactor isreheated and passed to a subsequent reactor and cools again. Thisprocess is repeated several times with a balance achieved between thenumber of reactors and the extent of conversion. The process stream iscycled several times through interreactor heaters to bring thetemperature back to a design inlet temperature. The process withmultiple reactors wherein the reactors are run until the temperature ofthe catalyst has dropped to a level where the conversion is too low tocontinue. The catalyst would then be withdrawn and reheated. This iscontrolled by controlling the flow of catalyst through the reactor. Theutilization of multiple reactors balances the reheating of the catalystand process stream with the length of contact time of the process streamat a high temperature, and allowances for the cooling due to theendothermicity. A second process (B) utilizes a large, back-mixed,reactor to generate a substantially uniform temperature in the reactor,with heat continuously added through the addition of fresh heatedcatalyst. The second process exposes the process stream to extendedperiods of high temperatures.

The use of multiple reactors requires the use of hot transfer lines, andextra heaters. This increases the lengths of time the process stream isat a high temperature, and can contribute to cracking, and reactionsbetween the process stream and hot metal in the equipment. This isovercome with a new approach to reduce the need to a single reactor andminimized the high temperature contact times, which in turn minimizesthermal cracking.

The new approach is to utilize a counter-current flow process forcontacting the paraffin with the catalyst. The new process utilizes alarger reactor, and allows for a larger gas flow rate, but has a lowercatalyst flux. The catalyst is controlled to flow downward whilecontacting the gas flowing upward in the reactor. The catalyst can bekept in the reactor for a longer residence time where the catalyst canbe allowed to cool more. There is an increase in yields while decreasingthe contact time between the process stream and the catalyst at therelatively high temperatures where undesired side reactions can occur.The catalyst residence time is determined by several factors, includingthe choice of catalyst.

This process is amenable to using a catalyst having a relatively shortlife before regeneration is required. For a short lived catalyst, suchas zirconia, the residence time can be between 10 seconds and 20minutes. For a catalyst having a longer life between regeneration, suchas a noble metal catalyst on a support, the residence time can be up toseveral days.

The process is for the dehydrogenation of propane and butane, and givesan improvement over current processes. The process, as shown in FIG. 2returns a regenerated catalyst stream 10 to a dehydrogenation reactor20. The catalyst flows downward through a catalyst bed 22 in the reactor20. A paraffinic feedstream 24 is passed to the reactor 20 and flows inan upward direction through the catalyst bed 22, thereby contacting thefeedstream and catalyst at dehydrogenation reaction conditions, togenerate a product stream 26 comprising olefins. Spent catalyst iscollected at the bottom of the reactor 20 and transferred through acatalyst transfer line 38 and passed to a regenerator 40. Theregenerator 40 regenerates the catalyst and returns the regeneratedcatalyst stream 10 to the reactor 20. The regenerated catalyst canundergo a stripping process with an inert gas to remove residualcombustion products from the regenerator. The inert gas can also beheated to maintain the catalyst temperature and to facilitate desorptionof adsorbed combustion products.

In the present invention, the dehydrogenation process includes areaction temperature between 400° C. and 800° C., with the temperaturegradient along the axial direction of the reactor. The reactor is at itshighest temperature at the top with the inlet of the regeneratedcatalyst and cools as the catalyst proceeds through the reactor. Thecatalyst feed is introduced to the reactor at a temperature of at least600° C., but less than 800° C. A preferred catalyst feed temperature isbetween 650° C. and 750° C., with a more preferred catalyst feedtemperature between 670° C. and 730° C.

The process is counter current, so the paraffinic feed, or processstream, is introduced at the bottom, or where the temperature is at itslowest. The paraffinic feed is introduced at a temperature no greaterthan 600° C., with the feed temperature of the process stream at least400° C. A preferred paraffinic feed stream temperature is between 450°C. and 550° C. and partial conversion is achieved at the relatively lowtemperature. A more preferred paraffinic feed temperature is between470° C. and 520° C.

This significantly reduces the amount of cracking of the hydrocarbons inthe process stream, by achieving some conversion at low temperatures.The equilibrium is shifts as the reaction proceeds, and to continue todrive the reaction the temperature needs to be increased to shift theequilibrium in a favorable direction. By passing the process streamcounter current to the catalyst, the equilibrium adjusts favorably asthe reaction proceeds and the process stream is exposed to an increasingtemperature as it passes through the reactor. The temperature profilesof the catalyst and feed are shown in FIG. 3, and are establishedthrough control of the flow of catalyst (C) and the process stream (P).The larger catalyst bed provides the heat for the reaction with thecatalyst cooling as it moves downward, and allowing for initialconversion in the feedstream at a lower temperature. While the profilesshown on FIG. 3 appear to have the process stream at a differenttemperature than the catalyst, the figure is showing the temperatures asthe material, either process stream or catalyst, enters a theoreticalstage. The points adjacent on the line would be the temperatures of thematerial either leaving one stage, or entering an adjacent stage. Thelines, (C) and (P), will in actuality more closely overlay each other,or the lines will be shifted relative to each other to the left or rightin the figure until a portion of the lines are overlaid. For example,the second point on the line (from the left) for the catalyst flow linewould have a close temperature to the first point of the line (from theleft) for the process stream. These would be the temperatures of thematerial leaving the upper theoretical stage.

The overall hot residence time is significantly reduced, and the processstream is exposed to high temperature only in the presence of catalystfor the short period of time as the process stream exits the catalystbed. The highest temperature exposure of the process stream is also forthe short contact time at the end of the process streams residence inthe reactor. The reaction conditions also include a reactor outletpressure between 20 kPa (absolute) and 400 kPa (absolute). Preferredreactor outlet pressures are from 105 kPa (absolute) to 300 kPa(absolute), with alternate operating conditions for the reactor outletpressure between 110 kPa (absolute) and 250 kPa (absolute), and between120 kPa (absolute) and 200 kPa (absolute).

The process further includes cooling the product stream 26. The productstream can be passed through a combined feed heat exchanger to preheatthe paraffinic feed stream, and cool the product stream. The process caninclude passing a portion of the cooled product stream to the upperregion of the reactor 20 through a quench line 36. The cooled portion ofthe product stream quenches the process stream as it comes off thecatalyst bed. The quenching inhibits further undesired side reactionsdue to the high temperatures, by rapidly cooling the process stream.

The process seeks to provide a substantially uniform distribution ofcatalyst across the catalyst bed for flowing down through the reactor.The catalyst returning from the regenerator 40 is passed to a catalystdistributor 30 to provide for the uniform distribution of regeneratedcatalyst over the bed. The process also seeks to maintain a uniformdistribution of the process stream flowing up through the catalyst bed.The paraffinic feed stream is passed to a feed distributor 32 to providea substantially uniform distribution of the feed stream across thecatalyst bed.

An aspect of the invention is to maintain a substantially uniformtemperature difference between the catalyst and the process stream asthe catalyst and process stream flow through the reactor. The catalystprovides the heat to drive the endothermic reaction, and as the reactionproceeds, the catalyst is cooled and flows downward. The reactants flowupward and are exposed to an increasing temperature to provide the heatfor the reaction, and to heat the process stream.

The process includes passing catalyst to the reactor at a catalyst inlettemperature of at least 600° C., and preferable at least 650° C. Thecatalyst is distributed over the top of the catalyst bed and flowsdownward through the reactor. The process includes passing a paraffinicfeed stream to a distributor disposed beneath the catalyst bed. The feedstream flows upward through the catalyst bed. The feed stream is passedto the reactor at a temperature of at least 50° C. below the catalystinlet temperature. To maintain the temperature difference between thecatalyst and the process stream, it is preferred that the paraffinicfeed stream inlet temperature is between 100° C. and 250° C. below thecatalyst inlet temperature, with a more preferred range of inlettemperature differences between 150° C. and 200° C. The control of thetemperature differences is partially dependent upon the catalyst flowrates through the reactor and the process stream flow rates through thereactor.

An aspect of the invention is the apparatus for the dehydrogenation ofhydrocarbons. The apparatus comprises a large reactor 20, as shown inFIG. 4. The reactor 20 can be a larger diameter reactor, up to 11meters, with a preferred range from 6 to 10 meters. This allows for ahigher gas flow rate, while having a lower catalyst flux within thereactor. Gas flow rates within the reactor are preferred to be from 0.4m/s to 1 m/s, with catalyst flux rates around 100,000 kg/m2/hr. Thecatalyst is distributed over the reactor bed with a catalyst distributor30, and the paraffin feedstream is distributed across the bottom of thecatalyst bed through a gas distributor 32. The reactor can includepacking grids 50, or other reactor internals for guiding the catalyst inthe downward direction and for reducing and limiting and axialback-mixing of the catalyst. Other reactor internals for this purposeinclude stripper trays such as used in an FCC unit, or gratings,provided the counter-current flow of catalyst and gas can be establishedwith limited or no axial backmixing.

The reactor 20 comprises a reactor shell having an upper region 52, acentral region 54 and a lower region 56. A catalyst inlet 62 is disposedin the upper region 52 of the shell, a catalyst outlet 64 disposed inthe lower region 56 of the shell, a process stream inlet 66 is disposedin the lower region 56 of the shell, and a product stream outlet 68 isdisposed in the upper region of the shell. The reactor 20 furtherincludes a set of reactor internals 50 disposed within the centralregion 54 of the shell, wherein the reactor internals 50 limit axialback-mixing of the catalyst.

The reactor design is for limiting the times the process stream isexposed to high temperatures. The reactor 20 includes a quench port 72disposed in the upper region 52 of the reactor shell. The quench port 72is for admitting a portion of the cooled product stream to rapidlyreduce the product temperature. The quench port can include adistributor for rapidly dispersing a cooled quench fluid into the upperregion 52 of the reactor.

The central region 54 of the reactor holds the reactor internals 50 andthe catalyst flowing through the reactor. The central region 54 providesfor a substantially uniform flow of catalyst, and has an axial length todiameter (L/D) ratio between 0.5 and 5, with a preferred length todiameter ratio between 0.6 and 2, and with a more preferred length todiameter ratio between 0.8 and 1.2. The axial length is the depth of thecatalyst bed, including reactor internals, for the reactor.

The central region 54 includes reactor internals that comprise astructured packing material. The structured packing includes a pluralityof ribbons, where each ribbon is angled, relative to the central axis ofthe reactor vessel 20. The ribbons refer to metal strips formed andangled relative to each other. The ribbons are arranged in arrays with aplurality of ribbons forming a layer. The packing can include multiplelayers of the ribbons. With multiple layers, each layer of ribbons canbe rotated around the central axis relative to each other to provide foraxial and azimuthal mixing of the catalyst. The neighboring layers, orneighboring packing units, are stacked to be in contact with neighboringunits. The units are preferably affixed to one another, or are stackedwith small spacers. Neighboring packing units are preferably rotatedrelative to each other around the central axis of the vessel. The amountof rotation is between 30° and 150° relative to a neighboring unit, andpreferably between 80° and 100°, and most preferably rotated at a 90°angle.

Each layer is attached to a neighboring layer, either through welding,or with mechanical items to affix the layers together, including rivets,screws, and other appropriate hardware. The ribbons can partiallyobstruct the passage of catalyst particles, and the angled portionprovides movement of the catalyst in either the radial direction,azimuthal direction, or a combination of both directions. For purposesof this invention, the movement in the radial direction refers to radialmovement toward or away from the central axis of the vessel;

the movement in the azimuthal direction is movement around the centralaxis of the vessel; and axial movement is movement parallel to thecentral axis of the vessel.

An enlarged view of two layers A, B of the structural packing 50 isshown in a perspective view in FIG. 5. Each ribbon 142 comprises bands154 configured in undulating peaks 162 and valleys 164. Each band 154includes a face 156 that obstructs passage of fluid and catalyst. Inthis embodiment, the bands 154 include laterals 155 arranged to providepeaks 162 at an upper landing 163 and valleys 164 at a lower landing165, but the peaks 162 and valleys 164 may be provided at the apex of ajoint of just two bands 154. The layers A, B each include paired ribbons142 a, 142 b. The lower landings 165 in upper ribbon 142 a meet theupper landings 163 of lower ribbon 142 b. A stabilizing strip 174 isdisposed between upper landing 163 and lower landing 165. If pairedribbons 142 a, 142 b are cut out of a common piece of metal, astabilizing strip 174 may be obviated. Ribbon 142 a is disposed at aphase that is 180° out of phase to the phase of paired ribbon 142 b.Other phase relationships may be used. Moreover, the axial spacing of aribbon 142 a is offset from the axial spacing of its paired ribbon 142b. Consequently, edges 158 of ribbon 142 a and edges 158 of ribbon 142 bmay be parallel and may define a plane therebetween. The edges 158 ofthe laterals 155 and landings 163, 165 in ribbon 142 a and the edges 158of the laterals 155 and landings 163, 165 in ribbon 142 b defineopenings 160 for the horizontal passage of fluid and catalyst. Edges oflaterals 155 and landings 163, 65 in alternating upper ribbons 142 a andalternating lower ribbons 142 b define openings 161 for the verticalpassage of fluid and catalyst. These openings 160, 161 are also definedby the faces 156 of the laterals 155 and upper and lower landings 163,165. Dimples 176 may be provided in the bands 154.

Although shown in laterals 155 near valleys 164, the dimples 176 may bedisposed in lower landings 165. It is also contemplated that edges 158of laterals 155 may be secured to each other in which case laterals 155would cross each other. Moreover, although the ribbons 142 arepreferably stacked horizontally in the central region 54, the ribbons142 may be arranged vertically in the central region 54. FIG. 5 showsvalleys 164 of lower ribbons 142 b in layer

A stacked on peaks 162 of upper ribbons 142 a in layer B. An alternatearrangement can have the landings 163, 165 oriented in a verticaldirection, or a direction parallel to the central axis of the vessel.

FIG. 6 is an enlarged partial perspective view of two segments 182, 184of one ribbon 142. Upper tabs 186 a, 186 b of adjacent segments 182,184, respectively, project from the standard strip 180 and may haveopposite configurations and be angular to each other. Lower tabs 188 a,188 b of adjacent segments 182, 184, respectively, project from standardstrip 180 and may have opposite configurations and be angular to eachother. Tie rods 198 extend through apertures 100 in standard strip 180to secure ribbons 142 in an array. The tie rod 198 may be welded to thestandard strip 180. Stabilizing strips 190 are seated in and secured totroughs 102 defined by upper tabs 186 a, 186 b and lower tabs 188 a, 188b of adjacent segments 182, 184.

The vanes, or ribbon sections that are angled with respect to thecentral axis of the vessel are between 5° and 60° from the central axis.Preferably, the vanes or ribbon sections are angled between 40° and 50°,and with a most preferred angle of about 45°.

While FIGS. 5 and 6 provide examples of the reactor internals, theinvention is not limited to these structures, but it is intended thatthe reactor internals comprise a configuration within the limitsspecified. For the purpose of this invention, the use of the term vanealso refers to the section of the ribbons, 155 or 186 or 188, that isangled with respect to the central axis of the vessel.

While it is preferred to have the structured packing units affixed toone another, due to design considerations, a small spacer can be placedbetween adjacent units. However, the spacing of the adjacent units is tobe less than 1 cm in the axial direction.

The reactor vessel 20 can includes a gas-solid separation system 74 forcapturing small catalyst particles that can become entrained with theproduct gas.

In a particular embodiment, a reactor for the dehydrogenation ofhydrocarbons comprises a reactor shell 20 having a generally cylindricalconfiguration and having an upper region 52, a central region 54 and alower region 56, a catalyst inlet 62 disposed in the upper region 52 ofthe shell, a catalyst outlet 64 disposed in the lower region 56 of theshell, a process stream inlet 66 disposed in the lower region 56 of theshell, and a process stream outlet 68 disposed in the upper region ofthe shell 52, wherein the upper region 52 of the reactor shell has agreater diameter than the central region 54 of the reactor shell. Thereactor includes a set of reactor internals 50 disposed within thecentral region 54 of the shell, wherein the reactor internals 50 are forguiding the flow of catalyst in a downward direction, while limitingaxial back-mixing of the catalyst. The reactor internals 50 are designedto limit the flow rate of the catalyst through the reactor and toprovide a means of distributing catalyst through the catalyst bed and toprovide for some mixing of the catalyst in the radial and azimuthaldirections. The reactor internals also prevent the creation and growthof gas bubbles within the bed. The reactor also includes a catalystdistributor 30 disposed in the upper region 52 of the shell and abovethe catalyst bed to provide a means to distribute catalyst over thecatalyst bed. The reactor also includes a process stream distributor 32disposed within the lower region 56 of the shell to distribute theprocess stream over the entire catalyst bed and to limit maldistributionof gas flow.

In this particular embodiment, the reactor includes a gas inlet port 72disposed in the upper region 52 of the shell for admitting a quenchfluid. The quench fluid provides for rapid cooling of the productstream. The reactor design also includes a gas-solid separation systemdisposed within the upper region 52 of the shell for recovering andreturning entrained catalyst particles to the catalyst bed.

The reactor internals 50 comprise vanes for moving the catalyst in aradial or azimuthal direction as the catalyst flows downward. The vanesare angled between 30° and 60° from the central axis of the vessel 20.

By controlling the flow of the catalyst, and having a larger amount ofcatalyst flowing in the downward axial direction, a temperature profilealong the axial direction can be established. This design promotes andeven fluidized bed density distribution throughout the reactor andinhibits the formation of zones having catalyst voids.

The dehydrogenation process comprises contacting a hydrocarbon streamwith a dehydrogenation catalyst at an elevated temperature.Dehydrogenation catalysts include noble metals on a support. Onecatalyst is platinum (Pt) on a support such as alumina. Another catalystis gallium oxide (Ga2O3) on a support, such as alumina, silica, activecarbon, or another refractory material. Other catalyst fordehydrogenation include chromic oxide (Cr2O3) on a support, molybdenum(Mo) on alumina (Al2O3), tin (Sn) promoted noble metals on a support,iron (Fe) and potassium (K) promoted chromic oxide, and copper (Cu) andcopper chromite. Supports include zeolites, alumina, silica-aluminas,zirconia, silica, magnesia, carbon and other refractory materials. Otherdehydrogenation catalyst include non-noble metal catalysts. In oneembodiment, a preferred catalyst is a non-noble metal catalyst such asnoble-metal free zirconia. The use of a noble-metal free catalyst saveson catalyst costs, and allows for flexible operation with a highcatalyst circulation rate.

One embodiment of the invention comprises passing catalyst from adehydrogenation catalyst distributor to the top of a catalyst bed in thedehydrogenation reactor. The catalyst flows downward under gravitythrough the reactor in the reaction section of the reactor. Ahydrocarbon feedstream comprising paraffinic compounds is passed to afeedstream distributor disposed at the bottom of the catalyst bed. Thehydrocarbon feedstream flows upward through the reactor and over thecatalyst in the catalyst bed, to generate a product stream comprisingolefins. The process stream is quenched in the upper part of the reactorvessel with a quench gas to reduce the temperature of the productstream. The quench gas can be any inert gas to lower the product streamtemperature to inhibit further thermal reactions.

The catalyst flows down through the catalyst bed through vanes to guidethe catalyst flow and to limit back-mixing of the catalyst in the axialdirection. Other guides, such as packing grids, or stripper packing canbe used within the catalyst bed.

The catalyst after passing through the catalyst bed is collected at thebottom of the reactor vessel. The catalyst is passed to a regenerationunit, and is preheated before passing into the regeneration unit. In analternative, and preferred method, the catalyst is regenerated usingadditional fuel. The additional fuel combusts and raises the catalysttemperature leaving the regenerator without having to preheat thecatalyst before returning the catalyst to the reactor.

The feedstream is heated before passing to the feedstream distributor.One embodiment for heating the feedstream comprises passing the productstream through a combined feed heat exchanger to preheat the feedstreamand to further cool the product stream. The preheated feedstream is thenheated to a desired feed inlet temperature between 450° C. and 550° C.

In one embodiment of the invention, the invention comprises a reactorfor use in the dehydrogenation of hydrocarbons. The reactor includes areactor shell, wherein the shell has a larger diameter than a normalreactor shell. The reactor shell includes an upper region, a centralregion, and a lower region. The reactor shell further includes acatalyst inlet disposed in the upper region, a catalyst outlet disposedin the lower region, a process stream inlet in the lower region, and aproduct stream outlet in the upper region. The reactor further includesin the central region of the reactor shell a set of reactor internals.The reactor internals comprise vanes for guiding the flow of catalyst ina downward direction and limit or restrict axial backmixing of thecatalyst. As the reactor shell will have a larger diameter, the catalystwill need to be distributed and the reactor will include a catalystdistributor disposed above the catalyst bed, in the upper region, touniformly distribute catalyst over the top of the catalyst bed. To limitlocal inhomogeneities, the process stream needs to be distributed overthe broader area beneath the catalyst bed, and the reactor includes aprocess stream distributor disposed beneath the catalyst bed, and in thelower region of the reactor shell.

The reactor can further include a gas inlet port disposed in the upperregion of the shell. The gas inlet port allows for the admission of aquench gas to cool the product stream after leaving the catalyst bed.The reactor shell can include a gas-solid separation system in the upperregion of the reactor shell to separate catalyst particles carried fromthe catalyst bed by the product stream.

The present invention provides for a much more favorable temperatureprofile in the axial direction, or the process stream flow direction,over current technologies that utilize a plurality of reactors withreactor interheaters. The reactors have falling temperature profileswith respect to the direction of hydrocarbon flow. This presentsimprovements over an adiabatic fixed bed reactor system operated in aswing mode where there is a shifting of stream between fixed bedreactors and the reactor beds are heated and regenerated in betweenreaction cycles. This also is an improvement over a fully back-mixedfluidized bed, as a back-mixed bed would have a flat temperatureprofile, and the process stream would have a longer exposure time to thehigh temperatures.

With this design a rising temperature profile, relative to the processstream flow direction, is created. This allows for partial conversion atlower temperature with the inlet stream and where the reactionequilibrium can still be favorable due to relatively low concentrationsof olefins. As the process stream encounters higher temperatures, theequilibrium shifts favorably, while limiting the high temperaturecontact time. A greater amount of catalyst can flow through the reactor,and the catalyst can be heated to supply the heat and flow more slowlythrough the reactor to sustain the endothermic dehydrogenation reaction.

While the invention has been described with what are presentlyconsidered the preferred embodiments, it is to be understood that theinvention is not limited to the disclosed embodiments, but it isintended to cover various modifications and equivalent arrangementsincluded within the scope of the appended claims.

1. A process for dehydrogenation of hydrocarbons comprising: flowing a catalyst in a catalyst bed, in a downward direction, through a dehydrogenation reactor, wherein the catalyst is added at a temperature of at least 650° C.; and flowing a paraffinic feedstream in an upward direction through the catalyst bed, thereby contacting the feedstream with the catalyst at dehydrogenation reaction conditions, thereby generating a product stream, wherein the feedstream is added at a temperature of at least 400° C. and less than 600° C., wherein the feedstream travels in an upward direction through the reactor,. providing for a substantially countercurrent flow of catalyst and paraffinic feedstream and wherein the paraffinic feedstream inlet temperature is between 100° C. and 250° C. below the catalyst inlet temperature.
 2. The process of claim 1 further comprising quenching the product stream.
 3. The process of claim 2 wherein the quenching of the product stream comprises adding a portion of a cooled product stream.
 4. The process of claim 1 further comprising passing the catalyst from a regenerator to a catalyst distributor disposed above the catalyst bed.
 5. The process of claim 1 further comprising collecting the catalyst at the bottom of the reactor to generate a return catalyst stream.
 6. The process of claim 5 further comprising passing the return catalyst stream to a catalyst regenerator.
 7. The process of claim 1 further comprising passing the paraffinic feedstream to a feedstream distributor disposed beneath the catalyst bed prior to flowing the feedstream over the catalyst bed.
 8. The process of claim 1 wherein the catalyst is passed to the top of the catalyst bed at a temperature between 650° C. and 800° C.
 9. The process of claim 8 wherein the catalyst is passed to the top of the catalyst bed at a temperature between 650° C. and 750° C.
 10. (canceled)
 11. The process of claim 1 wherein the paraffinic feed is at a temperature between 450° C. and 550° C.
 12. A process for dehydrogenation of hydrocarbons comprising: passing catalyst from a dehydrogenation reactor catalyst distributor to a dehydrogenation reactor, wherein the catalyst is admitted at a temp of at least 650° C.; flowing the catalyst in a catalyst bed, in a downward direction, through the reactor; collecting the catalyst at the bottom of the reactor to generate a return catalyst stream; passing a paraffinic feedstream to a feedstream distributor disposed beneath the catalyst bed, wherein the temp is at least 400° C. and no more than 600° C., and wherein the paraffinic feedstream inlet temperature is between 100° C. and 250° C. below the catalyst inlet temperature; flowing the feedstream in an upward direction through the catalyst bed, thereby contacting the feedstream with the catalyst, thereby generating a product stream, wherein the feedstream travels in an upward direction through the reactor providing for a substantially countercurrent flow of catalyst and paraffinic feedstream; and quenching the product stream.
 13. The process of claim 12 further comprising passing the product stream through a combined feed exchanger to preheat the paraffinic feedstream and cool the product stream.
 14. The process of claim 12 wherein the catalyst is returned from the regenerator at a temperature between 650° C. and 750° C.
 15. The process of claim 12 wherein the paraffinic feedstream is introduced to the reactor at a temperature between 400° C. and 550° C.
 16. The process of claim 12 wherein the catalyst is guided in the downward direction through the use of reactor internals to guide the catalyst in the downward flow.
 17. The process of claim 12 wherein the dehydrogenation reactor has a reactor product outlet pressure between 20 kPa (absolute) and 400 kPa (absolute).
 18. A process for dehydrogenation of hydrocarbons comprising: passing catalyst from a dehydrogenation reactor catalyst distributor to a dehydrogenation reactor, wherein the catalyst has an inlet temperature of at least 650° C.; flowing the catalyst in a catalyst bed, in a downward direction, through the reactor; collecting the catalyst at the bottom of the reactor to generate a return catalyst stream; regenerating and heating the catalyst in a catalyst regenerator to create a return catalyst stream; passing a paraffinic feedstream to a feedstream distributor disposed beneath the catalyst bed, wherein the temperature of the paraffinic feedstream is at least 100° C. below the catalyst inlet temperature; flowing the feedstream in an upward direction through the catalyst bed, thereby contacting the feedstream with the catalyst, thereby generating a product stream, wherein the feedstream travels in an upward direction through the reactor providing for a substantially countercurrent flow of catalyst and paraffinic feedstream; and quenching the product stream; wherein a substantially uniform temperature difference is maintained between the catalyst and the process stream as the catalyst and process stream flow through the dehydrogenation reactor.
 19. The process of claim 18 wherein the paraffinic feedstream has an inlet temperature between 100° C. and 250° C. below the inlet temperature of the catalyst.
 20. The process of claim 18 wherein the paraffinic feedstream has an inlet temperature between 150° C. and 200° C. below the inlet temperature of the catalyst. 